Hydrocarbon drying oil production



Dec. 27, 1955 s. E. JARos ETAL HYDROCARBON DRYING OIL PRODUCTION 2Sheets-Sheet Original Filed Dec.

s. E. JARos ETAL 2,728,801

HYDROCARBON DRYING OIL PRODUCTION Dec. 27, 1955 2 Sheets-Sheet 2Original Filed Deo. 25,. 1949 6561016 E. Soros Gather;-L-,Lecxsrzoveator's Qoberb Pf Lear @s CLtJtor'QezcS United StatesPatent Ctiice 2,728,801 Patented Dec. 27, 1955 HYDROCARBON DRYlNG DILPRODUCTION Stanley E. Jaros, Rahway, AnthonyH.` Gleason, Westfield, andRobert F. Leary, Cranford, N. J., assignors to Esso Researchl andEngineering Company, a `corporation of Delaware Original a piicationDecember 2.3,.; 1949 Serial No. 134,714.p Divided and this application.lime 16, 1952, Serial No. 293,772

5 Claims. (Cl. 260-666) This invention relates to a process for making adrying oil or varnish by polymerization of a diolefin such as butadienein the presence of sodium.r Thisis a. division of co-pending applicationSerial No. 134,714, led December 23, 1949. f

Drying oils of excellent quality can be made by polymerizing 100 partsof butadiene monomer, or particularly by copolymerizing about 75 to 85parts of butadiene-1,3 and about 25 to l5 parts of styrene at 25 to 95C., preferably at 65 to 85 C., in the presence of about 50i to 500 partsof an inert hydrocarbon ,diluent boiling between about and 250lC., orpreferably between 50 and 200 C., such as pentane, benzene, cyCIQheXane,naphtha or mineral spirits. Where low boiling materials are used, it isdesirable to operate under suficient pressure to maintain the charge inlliquidv phase, e. g. under pressures ranging from lk to atmospheres.,About 0.1

to 10 parts, preferably about 1 to 3 parts of nelydivided sodium metalis used as catalyst. Certain promoting agents, e. g. about 10 to 50parts of dioxaue and catalyst activators, e. g. 1 to 20 percent ofisopropyl alcohol based on the weight of sodium, are also preferablyadded to the reaction mixture to assure the production of a colorlessoil and to shorten the reaction time. Throughout this specification, allreferences to proportions ofy materials are expressed in parts byWeight, unless stated otherwise.

When the reaction has reached the desired conversion, which may be about50, S0 or preferably 100%, thecatalyst is destroyed and converted intoan easily tiltrable salt by addition of an essentially anhydrousmonobasic saturated C1 to C5 organic acid which is soluble inthehydrocarbon mixtt re. The acid, such as formic, acetic. or pentanoic, isadded directly to the crude reaction product and the resulting sodium salt is thereafter. separated from the crude product by iiltration. Thenthe crude filtrateA is fractionated toV remove the promoting andactivating agents as well as any unreacted monomer and promoter and alsoto adjust the concentration of the oily polymer in the product todesired level, which maybe between about 30-and 99% non-volatile matter.

Unlike natural` drying cils the resulting product is useful as a varnishwithout requiring any further. treatment for purposes of bodying-or thelike, but nevertheless itis often desirable to subject the product to alheating step for about 30 to 240 minutes at about 120 to- 180!1 C.whereby certain properties such as gloss of the .oil are improved. It isparticularly advantageous to lcarry out the. heating step in thepresence of a small amount of a polar compound such as maleic anhydride,acrylonitrile or thioglycolic acid.

When the original work was carried out in smally scale batches, itbecame apparent thatv a drying oil' or varnish of excellent quality wasobtained. At the same time, however, it became apparent that thetranslation .ofthe process into a large scale, commercial operationwould involve serious diiculties. For example, the handling anddestruction of relatively large amountstof metallic sodiumV represents'a manifest tireand explosion hazard. Secondly, whereasthe initialinduction period characteristic ofthe reaction was notr particularlydetrimental in smallr scale batch operation, since 4there it `took uponly a rather short time Vintervaly whereas the overall average reactionrate was quite rapid, this poisoningeiect equivalent to the inductionphenomenon became la probleml of the greatest magnitude lin large scalelor continuous-operation. v t

Heretofore, in a continuous process wherein the reagentsWerescontinuouslyfed to a one-stage equilibrium reactor and thepolymerized product continuously withdrawn therefrom after a certainresidence time, the unknown materials or conditions responsible for theinduction period poisoned the overall reactionrate to such an extent?that the required residence timebecame unecolnomically long; Theseparation of the various nonhydrocarbon constituents such asdioxanef,`a1cohol and aceticl acid from the crudeproduct alsorepresented an une-conotnicalstep when translated from experimental intolarge scale production, inasmuchl as normally'at least three corrosionresistant distillationy towers would be required. In-such anexperimental system the l'irst tower would serve-to separate thecrudeproduct into a bottom stream containingthe-oilyl polymer insolution in the hydrocarbon solvent and an overhead ystream containingexcess hydrocarbon solvent, unreacted monomer, ether promoter,alcohol-and acetic acid; this stream would' be fractionated in a secondtower to give an overhead stream containing the` ether promoterandalcohol anda bottom streamV containingthe hydrocarbon solvent and aceticacid; finally t-his bottomstreamy would. be fractionated in. a thirdytowertto give acetic acid inanA overhead stream and hydrocarbon solventyiny a bottom stream, it being extremely important to keep` any trace Vofacid from being recycled to the reaction-zone.

It is the principal objecty of 'the present invention to devisefanimproved processfor-the sodium polymerization. of .butadienoid dryingoils `on an industrial scale. This and. other objects will appear moreclearly from the subsequent description.

EigureHl-of the. accompanying drawing schematically illustrates thepreferred .equipment and flow sheet involved incarrying out theinvention;

Figure 2'. illustrates the preferredl type1 of polymerization reactor;

Figure.v 3 illustrates: an alternative typefof reactor; and

Figure 4 illustrates a polymerization system of still another type`Referring to. Fignrel, the reaction mixtureis preferably madeupinblending tank 1 into which butadiene, styrene, isopropyl;alcohol,`dioxane and straight run mineralspi'rits orother hydrocarbon diluent arevintroduced in the desired proportions. The reaction monomers andmake-up1quantitiesfof the alcoholareV introduced to tank 1 fromconventional storage facilities not shown and further,` amountsof thealcohol, dioxane-and the hydrocarbon diluentv are introduced. to tank 1through recycle line 10 which', leads from therecovery equipmentdescribedk later. After thematerials have been properly mixedy in tank 1by means of `stirrer2, astream` of the-resulting mixture is withdrawn'vthrough line 3to reactor 4. Suitable vreactor arrangementsxaredescribedv in detaill hereafter in reference-to Figures-2,. 3 vand 4.Priorto entering reactor 4 line Slmay be joined by li'ne 5 through whichsodium catalyst: isv introduced into reactor 4, preferably inthe formofr a slurry of finely` divided' catalyst particles in mineral spiritsorother suitable hydrocarbon liquid- Alternatively, catalyst.. line 5y mayenter the `reactor -separately; Thev average diameter of catalystparticles-useful in the present invention ranges betweenv about l1` and`l00 microns, preferably: between: 10 and 50 microns. Fur- Aor di-ethersof 2 to 8 carbons atoms, other than cyclic di-ethers of the meta-dioxanetype characterized by an O-C-O- ring structure, may be used likewisewhere not quite as good product quality or lower reaction rate can betolerated. Examples of such suitable ethers are diethyl ether,diisopropyl ether, diethyl acetal, ethylene glycol diethyl ether, furaneand the like.

After a reaction or residence time sulcient to allow for the desiredconversion of monomers, which time may range between about minutes and10 hours depending principally on size and design of reactor, reactiontemperature, catalyst concentration and particle size, the crudereaction product is withdrawn from reactor 4 through line 6 which leadsto diluent treating drum 7 which contains stirrer 9 and a suitablecooling device -such as internal coil 8 capable of removing any undueVheat of reaction that may be generated in this zone. In this drum thepartially spent sodium catalyst contained in the reaction mixture isutilized as a treating agent to Vremove or destroy all water and otherinhibiting impurities that are or may be present in the make-uphydrocarbon and `the other materials which are introduced into drum 7through line 11. Although the introduction of the diluent and othermaterials at this point somewhat increases the load on the fractionationequipment used in the recovery step described later, this or anequivalent pretreatment of the diluent with sodium is one of theessential elements of the present process in that it is surprisinglyeffective in improving the overall eiciency of the system, as opposed toa process wherein untreated make-up materials are fed directly to thereactor. This is particularly true as regards pre-treating thehydrocarbon diluent,- which constitutes the predominant fraction of themake-up streams in view of the fact that a substantial proportion of thediluent is eventually withdrawn in the final product. Conversely, inview of the relatively minute amount of dioxane required in the make-upstream solely to compensate for losses, it is feasible to introduce thelatter into the reaction without pre-treatment. The addition of thehydrocarbon diluent into drum 7 has the further advantage offacilitating subsequent ow and filtration of the crude polymerizationproduct which is quite viscous when cold, unless suitably diluted.

Water contained in the make-up diluent reacts with the sodium catalystin drum 7 and liberated hydrogen is withdrawn overhead through line 12.The crude product, diluted with an amount of make-up hydrocarbon solventapproximately compensating for the amount of solvent contained in thefinished product, is withdrawn from drum 7 through line 13 to anotherdrum 14 provided with cooling device and agitator 16. A slight excess ofanhydrous acetic acid is continuously introduced into drum 14'throughline 17 in order to react with the remaining sodium catalyst and therebyconvert it into sodium acetate which is an inert, easily filtrable salt.Hydrogen gas produced in this reaction is removed from acid treatingdrum 14 through line 18. From drum. 14 the acidifled crude product iscontinuously transferred through line 19 to drum 20 also preferablyprovided with a cooling device 21 and agitator 22. In drum 20 the smallamount of excess free acidfis neutralized with ammonia gas which isbubbled in through line 28. A suitable filter aid such as kieselguhr orfullers earth is also introduced into ammonia treating drum 20 so as tofacilitate subsequent filtration of the salts from the product. It isparticularly convenient to add the filter aid in the form of a thickslurry, e. g. as a dispersion in the same kind of hydrocarbon used asreaction diluent.

From drum 20 the crude product is withdrawn through line 24 and passed`through conventional filtration equipment such as a lter press or even acontinuous rotary vacuum lter 25. Eicient use of the latter is possibledue to the crystalline form of the filtrate obtained when operating inaccordance with the novel process. The ltered product is then run olf tostorage drum 26 which serves to smooth out any undue variations in owrate. From drum 26 the filtered product is fed continuously todistillation tower 27 which removes excess ammonia as an overhead stream28, concentrates a bottom stream 29 containing the oily polymer of thepreferred concentration of about 50 to 70% polymer in hydrocarbondiluent, and allows the withdrawal of a vapor side stream 47 whichcontains dioxane, any residual isopropyl alcohol as well as excesshydrocarbon diluent. The vapor stream 47 is condensed, passed to storagetank 48 and eventually recycled to blending tank 1 through line 49.

The ammonia stream 28 is recycled to treating tank 20 after beingsupplemented with fresh ammonia stream 23. Product stream 29 ispreferably passed through heat treating drum 30 where the drying oil iskept for about 30 to 120 minutes at a temperature of about 120 to 250C., preferably about 150 C., in order to improve the gloss of the oiland finally the product is passed through cooler 31 to storage tank 32.lf desired, a small amount, such as 0.01 to 1% (based on oily polymer)of' a polar compound such as maleic anhydride is also fed continuouslyto tank 30 through line 33 in order to further improve the varnishproperties of the product.

The storage tank is preferably provided with a steam coil or othersuitable heating device adapted to maintain the tank contents at atemperature of about 40 C. or higher in order to maintain the viscousproduct in a state sufliciently fluid for withdrawal.

Furthermore, all or a part of the product may be tapped o through line34 to be concentrated further to a concentration between and 99%,preferably between 95 and 98%. This can be done by any convenient means,as for example, by stripping the hydrocarbon solvent from the moredilute product by circulation through baffle drum 35, return line 36 andheater 37, at a temperature of about 150 C. to 200 C. under reducedpressure of about 0.5 to 2 lbs/sq. inch absolute. The concentratedproduct is finally withdrawn from the bottom of the tower through line38 and cooler 39 to storage tank 40, preferably maintained at about C.to maintain the viscosity of the product in a range permitting easypumping. The hydrocarbon solvent stripped overhead is passed throughline 41 and Water cooler 42 to separating drum 43 wherefrom thecondensed, dry hydrocarbon solvent is recycled either through lines 44,44a and 11 to solvent treating drum 7, or through lines 44, 44b and 10directly to blending tank 1. The vacuum necessary for the strippingoperation can be obtained readily by means of the illustratedarrangement consisting of aspirator line 45 connected to a suitablesteam jet ejector 46 which in turn is connected to condenser 47.

Alternatively, instead of concentrating the oily polymer in drum 35 asdescribed above, the drum may be replaced by a multiplate strippingtower operated under proper conditions, e. g. at a temperature of aboutC. and a pressure of 0.5 to 2 lbs/sq. inch absolute; the use of asuitable stripping gas such as methane improves the eliiciency of thetower.

The hydrocarbon diluent, alcohol and dioxane-containing vapor sidestream 47 withdrawn from principal fractionation tower 27 is recycledthrough storage tank 48 and return lines 49 and 10 to blending tank 1.

The choice of a proper reactor forms an important part of the novelprocess inasmuch as it has been discovered that with a proper system,substantially complete conversion of monomers may be achieved easily andquickly. This results in a greatly improved over-all economy inchemicals and process costs, since thereby any recovery of monomers fromthe product stream becomes superfluous and at the same time theotherwise inevitable loss of monomers due to escape or due toobjectionable polymerization in the recycle lines is eliminated.

In developing the present invention it was observed that, even with along nominal residence time, only very low conversion levels could beobtained when the reaction mixture was simply charged to and withdrawnfrom an agitated reaction tank in a continuous manner. It was alwaysrealized that continuous introduction of fresh feed into a one-stageequilibrium reactor theoretically can never be. driven to completeconversion, since this implies the absence of any fresh reactants in theproduct stream and therefore none in the reactor. However, even whenthis effect was taken into account, the conversion level was foundexperimentally to be surprisingly low in such a continuous equilibriumreactor. A possible explanation may be that the factors which areinherent in the type of feed used herein and which are responsible forthe initial induction period inV batch operations, the same factorspoison the continuous reaction to such an extent as to makepolymerization in a reaction tank continuously fed with fresh feedimpractical.

It has now been discovered that a high quality varnish' or drying oilcan be polymerized in a particularly ecient, continuous method by usinga reactor arrangement as shown in. Figure 2. The illustrated reactorconsists of an externally cooled or heated coil 241 connected to acooled reactor tank 242 provided with agitator 243 and nally anexternally cooled coil 244 connected to tank In this embodiment of theinvention the feed mixture from blending tank 1 (shown in Fig. l)together with a suitable amount of the sodium catalyst slurry from line5 (shown in Fig. l) are pumped continuously through heatedcoil 241 atsuch a rate that the feed passes through coil 241 in about 5 to 60minutes. During this passage the temperature of the reaction mixture isgradually raised from room temperature to about 40, or even 85 C.,preferably to a temperature between 50 and 65 C., and any feedimpurities presumably responsible for the induction period of thepolymerization reaction are destroyed, or the feed mixture is activatedby contact with the catalyst. From coil 241 the warm feed and catalystmixture is continuously introduced into agitated reaction tank 242 wheremost of the reaction, which is normally exothermic, takes place at atemperature between 65 and 85 C.

Tank 242 is provided with temperature control means (for heating orcooling) either by an external jacket or a heat exchanger constructedwithin the tank. For example, allowing for a nominal residence time ofabout 1 to hours at about 65 to 85 C., the monomer conversion in thistank will be between about 65 and 90 percent, whereupon this partiallypolymerized mixture is passed continuously through cooled coil 244 whoselength is such that passage of the reaction mixture therethrough takesabout 5 to 30 minutes, i. e. long enough to carry the reaction to 100%completion at a suitable temperature which is again maintainedpreferably between 65 and 85 C. It will be understood, of course, thatthe actual optimum length and size of the component parts of the reactordepend somewhat on composition and ow rate of the reaction mixture, theAreaction tempera ture, as well as the size relation between the severalcomponent parts of the reactor. It will be observed that each partserves a distinct function, namely, coil 241 serves principally tocontinuously activate or decontaminate the charge, though somepolymerization may occur even at this stage; reactor tank 242 serves tocontinuously polymerize the feed substantially as far as the chemicalequilibrium of the reaction mixture permits; and coil 244 serves tocontinuously polymerize the reaction to substantially completeconversion, which cannot be 0btained in tank 242. From coil 244 thecrude product passes through line 6 to the several treating drums andOther-finishing equipment shown -in Fig. 1 and described earlierherein.y

Instead of the ,reactorY arrangement shown in Fig. 2 a singlecontinuouscoil reactor may be used as illustrated inFig. 3, which` showsone of tive identical sections eventually joined in series to give areactor having the necessary length. Such a reactor can be readilyconstructedby joining together standard U.-shaped units 304, either toformk a` single` coil of the desired lengthl or to form smaller Sectionsas shown in Fig. 3, which are. finally connected in series to give thedesired length. For example, for a totalfeed rate of about 25 gallonsper minute (corresponding to ay monomer feed rate` of about 6 vgallonsof butadiene; and 1.5 gallons of styrene per minute) a reactionItemperature of about to 90 C. kand a 20- rninute residence time, andassuming a heatreaction of about 700 B. t. u. per pound of polymer, acoil` made of 11/2 inch standard pipe units should preferably have atotal length of about 4500 feet ASo'as.toprovideboth thenecessary heattransfer surface and the volume necessary to give the residence timerequired for complete conversion. If the reaction is carried out at alower temperature, e. g. at 50 C., the residence time must, be increasedcorrespondingly to about 60 minutes, in which event a coil reactor ofoptimum design may be constructed from 5-inch standard pipe, joinedtogether to a total length of about 1200 feet. The coil is providedwithr suitable external temperatureV control means such as the externaljackets 321 fed with a suitable heatingor coolingmedium from manifold322. If desired,v the first part of the coil may be heated, and the restof it cooled. Alternatively, the .entire reactor coil, or individualsections thereof,Y can be built within a large tank containing a`coolingmedium.

The use of individual sections joined in series isv preferable to aunitary coil reactor, since inthe former -case each section may have .aseparate thermostatic control andy hence the over-all reaction can becontrolled more readily and also more accurately. Moreover, whereas thepolymerization reaction in its early stages has a pronounced tendency torun away at elevated temperature levelsk due to the reactivity ofbutadiene monomer so thata relatively.. low initial polymerizationtemperature is preferable, the sectional arrangement permits one tooperate each successive section at progressively higher temperatures,with a corresponding saving in total reaction time and with increasedsafety as compared with a polymerization at a constant, comparabletemperature in aA unitary coil reactor. For instance, a particularlyadvantageous operation can be achieved by operating the rst coilsection, representing therst tenth or as much as theV rst third of totalcoil length, at a temperature between about 35 and 50 C., theintermediate coil section between 50 and 70 C., and the final coilsection between 70 and 90 C.

j Still another reactor arrangement suitable for carrying out thepresent invention by semi-continuous method is illustrated in Fig.. 4.In this embodimentline 3 leading from blending tank 1 (shown in Fig. 1)is connected to a manifold 401. At least three reactor tanks 402 areconnected to manifold 401 in parallel, each tank being provided withsuitable heating and cooling means 403, a stirrer 404, and valves 405,406 and safety valve or frangible disk 407. When operating withv thisarrangement of reactors valve 405 of the first tank is opened and valve406 is closed so that the tank is being gradually lled with freshreaction mixture and catalyst; inthe next tank,` previously charged withreaction mixture and catalyst, the charge is being batch polymerized toconversion while valve 406 is closed; at the same time the thirdl tankcontaining a previously polymerized batch is being gradually emptiedthrough valve 406 and line 6, valve 405 being closed and `vent 407opened to theatrnosphere or to a source of suitable linert gas such asmethane, nitrogen, or the like..V It will be` seen` that in this fashionthe overall process is maintained on a continuous production basis by asuitable switching of the `several batch reactors from a lling cycle toa polymerization cycle to a discharge cycle. A suitable duration of'each polymerization cycle as well as the filling and discharging cyclesmay be between about l and 24 hours depending on composition of feedmixture, catalyst concentration and particle size, total number and sizeof individual reactors, reaction temperature which may be between 50 and95 C., and so on. It must be pointed vout again, however, that althoughthe embodiment illustrated in Figure 4 is operative for the purposes ofthe present invention, it has numerous drawbacks as cornpared with thepreferred embodiment illustrated in Figure 2,v the principal drawbacksbeing the necessary duplication of large units of equipment, the lack oftrue continuity and consequent difficulty of controlling productuniformity and finally also a greater need for human supervision andmechanical control.

' A refinement further serving to increase the efficiency of the novelprocess and to raise the quality of the product involves feedingbutadiene and styrene monomers to different parts of the reaction zoneillustrated in Figure 2 in order to avoid the formation of polymerfractions having a high styrene content, which fractions tend to beformed in the early stages of the reaction. The formation of theseobjectionable polymer fractions can be avoided by not feeding thestyrene as shown in Figure 1, but instead styrene monomer is taken fromstorage directly to reactor 242 shown in Figure 2 so that no styrene ispresent in the initial stage of the reaction taking place in coil 2.41.`Where the single-coil reactor illustrated in Figure 3 is used, asimilarly favorable` result can be achieved by introducing styrene intothe reactor coil at an intermediate stage, e. g. after about one tenthto one third of the total coil length, instead of feeding it into thecoil together with the butadiene feed through line 3 as shown in Figure3. Finally, where batch reactors are used as illustrated in Figure 4,the same favorable effect may be achieved by withholding styrene monomerfrom the initial feed and instead introducing styrene into thepolymerization zone only at the end of the induction period, vor aboutto 30 minutes after the particular reactor has been put on stream.

The foregoing description is illustrative of the present invention.However, it will be apparent to those skilled in the art that numerousvariations and modiiications are possible without departing from thespirit of the invention or from the scope of the appended claims. Inreading the present specification and claims it should be understoodthat all quantities of materials are expressed in parts by weight unlessexpressly stated otherwise.

We claim:

l. A ycontinuous process for producing a hydrocarbon drying oil whichcomprises mixing 75 to 85 parts of butadiene, to 15 parts of styrene,200 to 300 parts of straight-run mineral spirits boiling between about150 and 200 C., 10 to 35 parts of p-dioxane, l to 3 parts of finelydispersed sodium metal and l0 to 20% of isopropyl alcohol based on theweight of sodium, heating the mixture to a reaction temperature between65 and 95 C. and continuously passing the heated mixture through anarrow tubular zone to activate the mixture, the length 'of the tubularzone being such that passage of the mixture therethrough takes about 5to 30 minutes,thereafter agitating the mixture in a closedpolymerization tank at reaction temperature until a monomer conversionof at least-'65% is reached, and continuously-withdrawing the partiallypolymerized mixture in linear ilow through a second tubular zonemaintained at reaction temperature, the length of said second tubularzone being such that complete conversion of monomers is obtained onpassage therethrough. 1

l 2. A continuous process for producing a'hydrocarbon drying oil whichcomprises mixing about 75- to 85 parts 8 of butadiene, 200 to 300 partsof straight run mineral spirits boiling between about 150 and 200 C., 10to 35 parts of p-dioxane, l to 3 parts of iinely dispersed sodium metal,and l0 to 20% of isopropyl alcohol based on the weight of sodium,heating the mixture to a temperature between 65 and 95 C. andcontinuously passing the heated mixture through a tubular zone for adistance equal to a time lapse of about 5 to 30 minutes, thereafterpassing the mitxure from the tubular zone to a closed polymerizationtank, continuously and separately feeding 25 to l5 parts of styrenemonomer to the polymerization tank, agitating the resulting mixture inthe polymerization tank at 65 to 95 C. for a nominal residence time ofabout l to lO hours, and continuously withdrawing the mixture from thepolymerization tank through a second tubular zone at 65 to 95 C. for atime lapse of about 5 to 30 minutes in order to complete thepolymerization of the entire monomer charge present in the mixture.

3. A continuous process for producing a hydrocarbon drying oil whichcomprises mixing 75 to 85 parts of butadiene, 25 to l5 parts of styrene,200 to 300 parts of straight run mineral spirits boiling between about150 and 200 C., l0 to 35 parts of an ether selected from the groupconsisting of mono and di-ethers of 2 to 8 carbon atoms and cyclicdi-ethers having 2 to 8 carbon atoms other than cyclic di-ethers of themetadioxane type characterized by an -O-C-G- ring structure, 0.1 to l0parts of nely dispersed sodium metal heating the mixture to the reactiontemperature between 65 and 95 C. and continuously passing the heatedmixture through a narrow tubular zone to activate the same, the lengthof the tubular zone being such that passage of the mixture therethroughtakes about 5 to 30 minutes, thereafter agitating the mixture in aclosed polymerization tank at reaction temperature until a monomerconversion of at least 65% is reached, and continuously withdrawing thepartially polymerized mixture in linear flow through a second tubularzone maintained at reaction temperature, the length of said secondtubular zone being such that complete conversion of monomers is obtainedon passage therethrough.

4. A continuous process for producing a hydrocarbon drying oil whichcomprises mixing about 75 to 85 parts of butadiene, 200 to 300 parts ofstraight run mineral spirits boiling between about 150 and 300 C., 10 to35 parts of an ether chosen from the group consistingof mono anddi-ethers having 2 to 8 carbon atoms and cyclic di-ethers having 2 to 8carbon atoms other than cyclic di-ethers of the metadioxane typecharacterized by an -O-C-O# ring structure, 0.1 to 10 parts of nelydispersed sodium metal, heating the mixture to a temperature between 65and 95 C. and continuously passing the heated mixture through a tubularzone for a distance equal to a time lapse of about 5 to 30 minutes,thereafter passing the mixture from the tubular zone to a closedpolymerization tank, continuously and separately feeding 25 to l5 partsof styrene monomer to the polymerization tank, agitating the resultingmixture in the polymerization tank at 65 to 95 C. for av nominalresidence time of about l to l0 hours, and continuously withdrawing themixture from the polymerization tank through a second tubular zone at 65to 95 C. for a time lapse of about 5 to 30 minutes inorder to completethe polymerization of the entire monomer charge present in the mixture.

5. A process for producing a polymeric hydrocarbon drying oil whichcomprises continuously mixing 75 to parts of butadiene monomer, 25 to l5parts of styrene monomer, 200 to 300 parts of a sodium-pretreated inerthydrocarbon diluent boiling between about and 200 C., about l0 to 35parts of p-dioxane, about l to l0 parts of finely divided metal sodiumcatalyst, and about l0 to 20% of alcohol based on the Weight of sodium,heating the mixture to a reaction temperature between 65 and 95 C. andcontinuously passing the heated mixture through a narrow tubular zone toactivate the mixture, the length of the tubular zone being such thatpassage of the mixture therethrough takes about 5 to 30 minutes,thereafter agitating the mixture in a closed polymerization tank atreaction temperature until a monomer conversion of at least 65% isreached, and continuously withdrawing the partially polymerized mixturein linear ow through a second tubular zone maintained at reactiontemperature, the length of said second tubular zone being such thatcomplete conversion of monomers is obtained on passage therethrough,withdrawing the polymerized sodium-containing mixture to a closedtreating zone, adding fresh hydrocarbon diluent in the treating zone tothe withdrawn mixture in order to make up for the quantity ofhydrocarbon diluent eventually Withdrawn in the product and whereby theresidual sodium reacts with and destroys water and other impuritiespresent in the added diluent, venting the said treating zone, passingthe resulting diluted mixture to an acid treaitng zone, adding an excessof glacial acetic acid to the diluted mixture in order to destroy allsodium metal, venting hydrogen from the acid treating zone, addingexcess ammonia gas to the acidified mixture in a neutralizing zone inorder to destroy al1 free acid, ltering the ammoniated mixture to removeall solids therefrom, fractionally distilling the filtered mixture,removing an overhead stream of ammonia from the distillation step andreturning the ammonia to the neutralizing zone, withdrawing hydrocarbonliquid, alcohol and dioxane in a vapor side stream from the distillationstep and recycling the side stream to the polymerization zone, andwithdrawing a concentrated bottom product stream from the distillationstep.

References Cited inthe file of this patent UNITED STATES PATENTS1,827,285 Ebert Oct. 13, 1931 1,832,450 Ebert Nov. 17, 1931 2,209,746Ebert .Tuly 30, 1940 2,264,811 Rothrock Dec. 2, 1941

1. A CONTINUOUS PROCESS FOR PRODUCING A HYDROCARBON DRYING OIL WHICHCOMPRISES MIXING 75 TO 85 PARTS OF BUTADIENE, 25 TO 15 PARTS OF STYRENE,200 TO 300 PARTS OF STRAIGHT-RUN MINERAL SPRITS BOILING BETWEEN ABOUT150 AND 200* C., 10 TO 35 PARTS OF P-DIOXANE, 1 TO 3 PARTS OF FINELYDISPERSED SODIUM METAL AND 10 TO 20% OF ISOPROPYL ALCOHOL BASED ON THEWEIGHT OF SODIUM, HEATING THE MIXTURE TO A REACTION TEMPERATURE BETWEEN65 AND 95* C. AND CONTINUOUSLY PASSING THE HEATED MIXTURE THROUGH ANARROW TUBULAR ZONE TO ACTIVATE THE MIXTURE, THE LENGTH OF THE TUBULARZONE BEING SUCH THAT PASSAGE OF THE MIXTURE THERETHROUGH TAKES ABOUT 5TO 30 MINUTES, THEREAFTER AGITATING THE MIXTURE IN A CLOSEDPOLYMERIZATION TANK AT REACTION TEMPERATURE UNTIL THE MONOMER CONVERSIONOF AT LEAST 65% IS REACHED, AND CONTINUOSLY WITHDRAWING THE PARTIALLYPOLYMERIZED MIXTURE IN LINEAR FLOW THROUGH A SECOND TUBULAR ZONEMAINTAINED AT REACTION TEMPERATURE, THE LENGTH OF SAID SECOND TUBULARZONE BEING SUCH THAT COMPLETE CONVERSION OF SAID TUBULAR ZONE BEING SUCHTHAT SAGE THERETHROUGH.